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Ind. Eng. Chem. Res. 2010, 49, 682–689
Conceptual Evaluation of Integrated Process Configurations for the Recovery of L-Phenylalanine Product Crystals during Fermentation Maria C. Cuellar,† Adrie J. J. Straathof,*,† Emile J. A. X. van de Sandt,‡ Joseph J. Heijnen,† and Luuk A. M. van der Wielen† Department of Biotechnology, Delft UniVersity of Technology, Julianalaan 67 2628BC, Delft, The Netherlands, and DSM Biotechnology Center, P.O. Box 1, 2600MA, Delft, The Netherlands
The production of the amino acid L-phenylalanine (Phe) by fermentation is in many cases limited by the tight regulation of the microbial pathway leading to Phe synthesis. One way of circumventing feedback repression is to remove the product from the vicinity of the micro-organism as soon as it is being formed. In the case of Phe, techniques like adsorption and extraction have been applied with this aim. In these processes, however, additional recovery steps are required in order to obtain the product as anhydrate crystals, which is its commercial form. In this work we evaluated conceptually the recovery of Phe anhydrate crystals during a fermentation process. The product recovery consisted of water removal by reverse osmosis, crystallization, and recycle of the mother liquor either to the fermenter, to the water removal unit or to both. By maintaining the Phe mass fraction in the fermenter at about 17 g kg-1 the fermentation productivity increased according to the calculations from 0.66 g kg-1 h-1, without product removal, to 1.07 g kg-1 h-1, with product removal, where about 70% of the produced Phe was directly recovered as anhydrate crystals. By means of a simplified economic model it was shown that the membranes required for cell retention and water removal have much more impact on the economic performance of the process than the consumption of raw materials, and therefore, the favored recycling option for the mother liquor is to the water removal unit. Introduction L-Phenylalanine (Phe) is one of the most important commercially produced amino acids. Besides being used as a feed and food supplement, Phe is used as a building block for the synthesis of the artificial sweetener aspartame. The world consumption of Phe was estimated at about 14 000 t in 20021 with an increasing market due to the high demand for lowcaloric foods and beverages.2 Production of Phe at large-scale can be achieved by several chemical, biocatalytic, and fermentation methods3 from which the last two dominate the market.2 In the production by fermentation, however, the metabolic pathway leading to the production of Phe (i.e., the aromatic amino acids pathway) is tightly controlled. The presence of Phe in the fermentation broth at levels higher than 3 g L-1 significantly reduces the activity of many enzymes involved in Phe synthesis.4 With the rise of metabolic engineering, over the last decades a large number of strains has been developed for the (over)production of Phe. Many of these strains belong to the genera Corynebacterium and BreVibacterium, well-known in the large scale production of amino acids, but much research has also been devoted to the faster-growing Escherichia coli. Metabolic engineering has not been the only tool for fermentation process improvement. Process engineering, for example, has been useful in several processes for control over the production of byproducts like acetate in fermentations with E. coli.5,6 Furthermore, to increase the fermentation productivity with strains for which the regulation has not been fully alleviated, product removal during fermentation (also referred to as in situ product removal, in situ product recovery or ISPR) has been implemented. Kusunose7 used uncharged polymeric beads to remove Phe during a fermentation with BreVibacterium
* To whom correspondence should be addressed. E-mail: A.J.J.
[email protected]. Tel.: +31-15-278 2330. Fax: +31-15-278 2355. † Delft University of Technology. ‡ DSM Biotechnology Center.
lactofermentum, while a process integrating reactive extraction, either in hollow fiber membranes4,8 or in centrifuges,9 to a fermentation with recombinant E. coli was extensively studied up to pilot-plant scale. All these processes resulted in increased fermentation productivity and product yield on substrate (glucose) when compared to a reference fermentation without product removal. Next to this, the processes with reactive extraction allowed the online concentration of Phe, which was beneficial for its subsequent crystallization.9 Within the frame of product removal during fermentation, yet another alternative is to recover the product by crystallization. This alternative is particularly attractive for products that are commercialized in crystal form. When the product concentration is sufficiently high, crystallization might take place in the fermenter; otherwise, crystallization might be carried out outside the fermenter as described by Buque-Taboada et al.10 In the case of Phe, two crystal forms are known: a flake-like anhydrate and a needle-like monohydrate. Anhydrate is the preferred commercial form because of its ease of downstream processing.11 In water, anhydrate is thermodynamically stable above 37 °C.11 Monohydrate, however, has been reported to be the first form to appear during spontaneous crystallization in water12 and during fermentation, resulting in viscosity increase and overfoaming.13 This has also been observed with similar compounds such as L-tyrosine.14 Thus, the implementation of product recovery during Phe fermentation should aim at the production of anhydrate crystals. Igarashi et al.13 described a fermentation process with B. lactofermentum in which anhydrate crystals were produced in the fermenter. After the fermentation, a concentration and four recrystallization steps were required in order to recover about 60% of the total Phe produced. In many Phe fermentations, however, product inhibition takes place well before supersaturation is reached. In such cases, crystallization outside the fermenter is a better alternative. In this work, we evaluate conceptually the production of Phe anhydrate crystals during fermentation. The fermentation process
10.1021/ie901007g 2010 American Chemical Society Published on Web 11/20/2009
Ind. Eng. Chem. Res., Vol. 49, No. 2, 2010
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Figure 1. Schematic representation of a fermentation operating with cell recycle and product recovery by crystallization. Product recovery is achieved by first concentrating the cell-free solution from the fermenter in a buffer vessel coupled to a membrane filtration unit for water removal. The concentrated solution is fed to a crystallizer operating with crystal retention.
using a recombinant E. coli strain as described by Takors8 is used as starting point. The product recovery is performed in two steps: concentration and crystallization. The objective is 2-fold: to evaluate whether product removal by crystallization can increase the fermentation productivity, and to define the most suited process configuration on the basis of the costs associated to it. This approach is useful in the early stages of process design and should be applicable to other processes involving product recovery by crystallization coupled to fermentation. Theory and Model Development Figure 1 gives a schematic representation of a fermentation operating with cell retention and product recovery by crystallization. The product recovery comprises a water removal step by membrane filtration and crystallization in a separate vessel with crystal retention. The flows between the membrane units and their feed vessels are very large compared to the other flows. Therefore, the vessels and the membrane units are modeled as one well-mixed unit (dashed boxes in the figure). From this point on these well-mixed units will be referred to as fermenter and water removal unit. The fermentation is based on the process described by Takors.8 In this process glucose is used as the sole carbon source. Additionally, L-tyrosine (Tyr) is fed to the process because a Tyr-auxotrophic, recombinant E. coli is used. The cell growth is limited by Tyr, and the glucose mass fraction is controlled at around 5 g kg-1 so that the production of the byproduct acetate is minimized without limiting the Phe production capacity.4,8 Phe production is induced using IPTG. In the fermenter, the process starts with cell batch growth; when Tyr is depleted and glucose is reduced to about 5 g kg-1 the Tyr and glucose feeds are started in order to extend the cell growth up to about 30 g kg-1 of cell mass. At this point the cell-free solution is transferred to the water removal unit for concentration, the flow to and from the crystallizer is started, and anhydrate seed crystals15 are added to the solution in order to start the crystallization. The permeate obtained during water
Figure 2. Schematic representation of concept A and concept B. The boxes correspond to the unit operations involved: fermentation with cell retention (1), water removal by membrane filtration (2), and crystallization with crystal retention (3). See text for further details.
removal, which also contains Phe, may be recycled to the fermenter (see Figure 2A). Furthermore, as described in Buque-Taboada et al.10 the solution from the crystallizer (i.e., the mother liquor) still contains nutrients and dissolved product, and therefore it can be recycled to the fermenter or to the water removal unit. All these recycling possibilities are depicted in
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Figure 2B. Note that the only purge in such system is Φout. Other purge streams are possible, but are not considered in this work. In this model the system is operated as a repeated fed-batch of 50 h, which is the same time span of the fermentation process described by Takors.8 Consequently, the system is assumed to operate with prefilled buffer vessel and crystallizer; that is, the time and mass required to start up the operation of the product recovery loop is not taken into account. 1. Fermentation Equations. Similar to the approach described by Takors,8 in this model only main components are taken into account: cells, Phe, Tyr, glucose, and water. It is assumed that the specific cell growth rate µ is only dependent on Tyr, since glucose is always supplied in excess: µ ) µmax
aq CTyr,f aq CTyr,f + KTyr
qPhe ) qPhe,max
aq KI,Phe + CPhe,f
qTyr ) qGlu )
+ qm,Tyr
YX/Tyr
µ YX/Glu
+
qPhe YPhe/Glu
+ qm,Glu
(3) (4)
For a Tyr-limited fed-batch: aq dMTyr,f ≈0 dt
(5)
aq ΦTyrCTyr,feed ) qTyrMX,f
(6)
It is assumed that the cells are fully retained by the cell retention membrane. The cell mass balance becomes dMX,f aq ) ΦTyrCTyr,feed YX/Tyr - qm,TyrYX/TyrMX,f dt
(7)
The glucose level in the fermenter is controlled around a set -1 point Caq Glu,f,set ) 5 g kg . Before the product recovery is started the variable glucose feed rate can be calculated as ΦGlu )
qGlu aq CGlu,feed
-
aq CGlu,f,set
dMX,f dt
(8)
Membrane Equations. Following the same approach described in Cuellar et al.,15 the mass transfer across the membrane is described by a solution/diffusion model. The water flux Jw depends on the water permeability of the membrane, Aw, and the net transmembrane pressure: Jw ) Aw(∆PTM - ∆ΠTM)
(9)
The transmembrane pressure ∆PTM and the transmembrane osmotic pressure ∆ΠTM are calculated as the difference between the pressure at the membrane (Pm or Πm, respectively) and the pressure at the permeate (Pm or Πp, respectively). The osmotic pressure Π is calculated from the Van’t Hoff equation:
)
(10)
( )
aq aq aq ) BPhe(CPhe,b - CPhe,p ) exp JPhe ) JwCPhe,p
Jw kPhe
(11)
This is the same for the glucose and Tyr fluxes. It is assumed that the mass transfer coefficients for glucose and Tyr are comparable to those of Phe. The selectivity of the membrane is conventionally measured as the retention (also called rejection) coefficient: RPhe ) 1 -
aq CPhe,p
(12)
aq CPhe,b
Combining eq 11 and 12 and solving for RPhe: RPhe )
(2)
where KI,Phe is the inhibition constant. The specific consumption rate of Tyr takes into account the requirements for cell growth and maintenance; the glucose counterpart additionally takes product formation into account: µ
(
aq aq aq CGlu CTyr CPhe + + WPhe WGlu WTyr
The Phe flux JPhe depends on the mass transfer coefficients in the membrane, BPhe, and in the polarized layer, kPhe:
(1)
Phe production is growth-decoupled but the cell mass-specific production rate qPhe is inhibited by high product levels: KI,Phe
Π ) RTF
Jw Jw BPhe exp + Jw kPhe
( )
(13)
3. Crystallization Equations. It is assumed that only Phe crystallizes and that glucose and Tyr do not exert any effect on the crystallization of Phe. In the same way as described in Cuellar et al.,15 for a system with crystal retention: s pPhe,c
)
s 3(MPhe,c )2/3
(
s aq (MPhe,c,0 )1/3 MPhe,c kg aq aq -1 L0 Mc CPhe,c,eq
)
g
(14)
s is the mass of seeds added to the crystallizer, where MPhe,c,0 is the solubility of Phe at the crystallization temperature, Caq Phe,c,eq L0 is the average size of the seeds, and kg and g are the crystal growth kinetics constants. 4. Simulation of the Total Process. The equations described above were incorporated in the mass balances for the fermenter, the water removal unit, and the crystallizer. These mass balance equations were solved simultaneously using a second order Runge-Kutta routine16 in Matlab 7.3 (Mathworks, USA). This model is used to describe the fermentation with and without product removal. In the fermentation with product removal, the model allows the evaluation of the accumulation of some components due to the recycles. In this way, the reduction in glucose consumption in some configurations (see also the section Simulations at the Pilot Plant Scale) can be aq aq and CGlu,c can be evaluated accounted for. Furthermore, CTyr,c in time in order to assess whether accumulation might result in crystallization of glucose or Tyr. Finally, Caq Phe,b can be evaluated in time in order to assess whether the metastable limit is passed, risking the formation of the wrong crystal form (monohydrate). 5. Steady-State Situation. As described in Cuellar et al.,15 for the product recovery section a constant crystal production aq rate can be achieved by keeping CPhe,c constant at steady state. Given
aq CPhe,c )
aq MPhe,c
Maq c
(15)
The mass balance equations for the crystallizer are dMaq c s ) Φc,in - Φc,out - pPhe,c dt
(16)
Ind. Eng. Chem. Res., Vol. 49, No. 2, 2010 aq dMPhe,c
Table 2. Total Manufacturing Cost (TMC) Calculation aq aq s ) Φc,inCPhe,b - Φc,outCPhe,c - pPhe,c
(17)
dt At steady state eqs 16 and 17 are equal to 0. Solving for Φc,in: s Φc,in ) pPhe,c
aq 1 - CPhe,c aq CPhe,b
-
(18)
aq CPhe,c
aq In the water removal unit CPhe,b and Mb should also remain constant. Hence,
Φf ) Φp + Φc,in - Φr,3 aq ΦfCPhe,f
)
aq ΦpCPhe,p
+
aq Φc,inCPhe,b
-
(19) aq Φr,3CPhe,c
(20)
where Φr,3 ) 0 in concept A. The permeate flow can also be expressed as Φp ) JwARO
(21)
where ARO is the membrane area. The concentration factor in the buffer vessel is defined as F)
aq CPhe,b
(22)
aq CPhe,f
Substitution of eqs 12-13 and 19-22 allows the calculation aq of the required ARO for a given CPhe,c and Jw: ARO )
(
(
))
aq CPhe,c 1 Φc,in(F - 1) - Φr,3 aq - 1 Jw(1 - (1 - R)F) CPhe,f (23)
where Φr,3 ) 0 in concept A. Once the product recovery and recycles are operational, the overall system may reach a pseudosteady-state. For concept A thus: ΦGlu + ΦTyr ) Φout + Φc,out ΦGlu + ΦTyr ) Φout
cost
variable costs (VC) fixed costs: maintenance operating labor (OL) laboratory costs supervision plant overhead capital charges patents and royalties local taxes insurance total general expenses
calculated 0.05 0.10 0.20 0.20 0.50 0.15 0.01 0.02 0.01 0.43 0.25
FCI FCI OL ) 0.02 FCI OL ) 0.02 FCI OL ) 0.05 FCI FCI FCI FCI FCI FCI × (0.43 FCI + VC)
Table 3. Parameters Used in the Model parameter µmax KTyr YX/Tyr qm,Tyr qPhe,max KI,Phe YPhe/Glu YX/Glu qm,Glu Aw BPhe kPhe kg g
value 0.321 0.00152 35.29 0.00245 0.0756 20 0.138 0.51 0.04 5.6 × 10-6 3.32 × 10-5 0.0099 2.37 × 10-6 1.8
unit
reference
h g kg-1 g g-1 g g-1 h-1 g g-1 h-1 g kg-1 g g-1 g g-1 g g-1 h-1 s m-1 kg m-2 s-1 kg m-2 s-1 m s-1 -
8 8 8 8 8 8 4 20 20 15 15 15 15 15
-1
estimated as fractions of the fixed capital investment (FCI) as shown in Table 2. The TMC calculation can be summarized as
(24)
TMC ) 1.25(0.43 FCI + VC) ) 0.54 FCI + 1.25 VC (26)
(25)
Subsequently the key economic numbers can be calculated. In this work, the net present value (NPV) is used for comparing the economic performance of the different process configurations:
and for concept B, s pPhe,c
item
685
17
aq F,CPhe,c
and given and Pm, this steady-state For a required model allows the calculation of Φf, Φc,in and ARO. 5. Simplified Economic Model. A simple economic model allows for a quick comparison of process configurations based on their impact on variable cost (VC), purchased equipment cost (PEC), and revenues. The total capital investment (TCI) can be calculated as a function of PEC as shown in Table 1. The total manufacturing cost (TMC) is calculated as the sum of VC, fixed costs, and general expenses. The fixed costs and the general expenses are Table 1. Total Capital Investment (TCI) Calculation.17,18 item
cost
purchased equipment costs (PEC) equipment installation instrumentation piping electrical buildings and storages site development direct plant costs (DPC) design and engineering contractors’ fee contingency indirect plant costs (IPC) fixed capital investment (FCI) working capital start up total capital investment (TCI)
PEC 0.40 PEC 0.20 PEC 0.70 PEC 0.10 PEC 0.45 PEC 0.05 PEC 2.90 PEC 0.30 DPC 0.05 DPC 0.10 DPC 0.45 DPC ) 1.30 PEC DPC + IPC ) 4.20 PEC 0.10 FCI 0.08 FCI 1.18 FCI ) 4.96 PEC
NPV) annual cash flow × (1 + interest rate)-i The following assumptions are made: (1) TCI is invested in year i ) 0. (2) Production is full scale in years 1-9. (3) FCI is linearly depreciated in these 10 years (depreciation ) 0.1 × FCI). (4) The income tax rate is 30%. (5) NPV after these 10 years should be maximized for an interest rate of 10%. The cash flow is constant for years i ) 1-9. Cumulating (1 + 0.10)-i for years i ) 1-9 gives 5.76. Subsequently, NPV ) 5.76 × annual cash flow - TCI (27) The annual cash flow for years i ) 1-9 is obtained from the annual gross profit minus taxable profit: annual cash flow ) (revenues - TMC) (revenues - TMC - depreciation) × 0.3 (28) Substituting eqs 26 and 28 in eq 27 and expressing FCI as a function of PEC (Table 1): NPV ) 4 revenues - 5 VC - 13 PEC (29) Simulations at Pilot Plant Scale. For the simulation of the fermentation, the pilot plant fermenter as described in Takors8 is assumed. The fermenter had a broth capacity of about 300 kg and it was started with 123 kg.4 Table 3 shows a summary of the parameters used in the model. The fermentation model by Takors8 does not include kinetic parameters for glucose;
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Table 4. Recycle Configurations Evaluated in This Studya
a
configuration
Φr,1
B1 B2 B3 B4 B5
variable variable variable variable variable
Table 5. Simulation Resultsa
Φr,2
Φr,3
Φc,out /4Φc,out 1 /2Φc,out 1 /4Φc,out 0
0 1 /4Φc,out 1 /2Φc,out 3 /4Φc,out Φc,out
3
See Figure 2B and text for further details.
therefore, it is assumed that glucose consumption for cell growth and maintenance is comparable to that of a wild-type strain, whereas the Phe yield on glucose is taken from previous studies with the same strain used by Takors.8 For water removal, a reverse osmosis membrane as described in Cuellar et al.15 is considered. In that study, a flat-sheet membrane (reference SE, GE Osmonics Inc., Minnetonka, MN) with an active area of 0.014 m2 and maximal operating pressure of 4000 kPa was evaluated in a Sepa CF II Membrane Cell System (GE Osmonics Inc., Minnetonka, MN) operating in cross-flow mode. According to the supplier, such system resembles the hydrodynamics of a large scale membrane module. Therefore, in this model the parameters determined by Cuellar et al.15 are used (see Table 3). To increase the fermentation productivity by means of product aq should remain low. For a given CX,f the maximal removal CPhe,f aq ) 0 g kg-1, fermentation productivity is achieved when CPhe,f aq see eq 2. The practically feasible value of CPhe,f depends on the capacity of the product removal technique. Independent of the technique, however, the product should be removed at the same rate as it is produced;10 otherwise the product will continue to accumulate in the fermenter and the production rate will decrease. A high Caq Phe,f, on the other hand, simplifies the product aq recovery.19 The system under consideration is targeted at CPhe,f e KI,Phe. With the fermentation model only, it was found that aq reaches 20 g kg-1 after about 25 h of fermentation CPhe,f (simulation results not shown). Because of product inhibition, at that point the cell mass-specific Phe production rate qPhe is 0.0378 g g-1 h-1 (50% of qPhe,max). The cell mass in the fermenter MX,f is about 6220 g, which translates in a Phe production rate of 235 g h-1 (653 × 10-7 kg s-1). The product recovery section should thus be designed in such s equals this production rate, and that the target a way that pPhe,c anhydrate crystals are produced. In a similar way as described in Cuellar et al.,15 the water removal unit is operated at 50 °C, the Phe mass fraction in this unit is maintained below the aq e 47 g kg-1)15 and crystallization is metastability limit (CPhe,b aq is 41 and performed at 45 °C. At these temperatures, CPhe,c,eq 38 g kg-1, respectively.15 The recycle configuration shown in Figure 2A has a limited aq applicability. From eq 18 it can be seen that, for a given CPhe,b aq s and CPhe,c, a higher Φc,in results in higher pPhe,c. In this configuration, the highest value for Φc,in is achieved when Φp is fully recycled; Φc,in equals then the feed flow rates (3.74 kg aq ) 47 and h-1 at about 25 h of fermentation). Assuming CPhe,b aq -1 s CPhe,c ) 39 g kg and following eq 18, pPhe,c ) 31 g h-1 (0.9 aq aq and CPhe,c are expressed × 10-7 kg s-1; note that in eq 18 CPhe,b s is much lower than in g g-1). Thus, the obtained value of pPhe,c s recycles need required. Therefore, to achieve the required pPhe,c to be implemented as shown in Figure 2B. Note that for other systems with a wider window of operation and lower solubility, configuration A might be sufficient and even preferred because of its simplicity. Table 4 shows the recycle ratios according to Figure 2B evaluated in this study. In all cases, the operating pressure was
configuration B1 B2 B3 B4 B5
aq b aq b aq b Φf ARO CTyr,c CGlu,c CPhe,b Phe recoveryc (kg s-1) (m2) (g kg-1) (g kg-1) (g kg-1) lossd
2.9 2.2 1.6 1.0 0.4
× × × × ×
10-2 10-2 10-2 10-2 10-2
2.67 2.15 1.62 1.10 0.75
0.017 0.018 0.019 0.021 0.038
12.46 12.99 14.00 16.86 48.05
45.58 45.60 45.48 45.23 43.70
71.4 72.0 71.3 70.3 69.9
0.34 0.37 0.38 0.40 0.48
a
See text for further details. b At the end of the batch. c Percentage of Phe crystal mass relative to the total Phe mass produced. d Percentage of Phe mass lost in Φout relative to the total Phe mass produced.
Figure 3. Simulation profiles for configuration B3: (Top) mass fraction profiles in the fermenter; (bottom) Phe mass fraction profiles in each unit. The vertical dotted line shows the start of the product recovery.
set to 3000 kPa. Φc,out is fully recycled, while Φr,1 varies for each configuration as required to maintain the fermentation mass constant. Φf,Φc,in and ARO were determined following eqs 18, 23, and 25. In this way the required psPhe,c resulted in Φc,in ) 1.1 × 10-2 kg s-1 for all configurations. With these values the mass balance equations for the complete system were solved using Matlab 7.3. Table 5 shows some of the results obtained for each configuration and Figure 3 shows the profiles for configuration B3. These profiles were comparable for all configurations. In all cases, the fermenter mass Mf and the cell mass were kept constant at 186 and 6 kg, respectively. The Phe mass fraction aq was about 17 g kg-1 and the total Phe in the fermenter CPhe,f produced after 50 h averaged 9960 g. This resulted in an overall fermenter mass-specific productivity of 1.07 g kg-1 h-1, which is considerably higher than that of a fermentation without product removal (0.66 g kg-1 h-1; total Phe produced ) 9330 g and fermenter mass ) 284 kg). From Table 5 it can be seen that a larger Φr,2 results in a larger Φf and ARO. Also shown in Table 5 is the percentage of Phe lost in Φout; a larger Φf leads to larger Φr,1 and consequently,
Ind. Eng. Chem. Res., Vol. 49, No. 2, 2010 Table 6. Prices Used in the Simplified Cost Model. Reference Year: 2001 unit variable costs (VC′) glucose tyrosine replacement UF membrane replacement RO membrane purchased equipment costs (PEC′) membrane unit (spiral wound) revenue phenylalanine
price per unit ($)
Table 8. NPV′ Calculation for All Process Configurations VC′
ref
kg kg m2 m2
0.31 34 37.5b 20c
USDAa 23 22 22
m2
400
22
kg
20
23
687
a
configuration
raw materials
membrane replacement
PEC′
revenue
NPV′
B1 B2 B3 B4 B5
7763 7905 8051 8192 8405
208 163 119 75 35
2715 2143 1572 1002 513
22862 23017 22768 22358 22223
16295 23871 29785 35078 40021
a
Prices in dollars.
a United States Department of Agriculture. Economic research service: wholesale list price for glucose syrup, midwest markets, http:// www.ers.usda.gov/briefing/sugar/data.htm. b Polysulfone, spiral wound. c Thin film, spiral wound.
Table 7. Area Required for Cell Retention (AUF), Glucose Consumption, and Phe Production per Yeara configuration
AUF (m2)
glucose (kg/a)
Phe (kg/a)
B1 B2 B3 B4 B5
4.12 3.21 2.31 1.40 0.53
12723 13182 13651 14106 14794
1143 1151 1138 1118 1111
a
Tyr consumption was 112 kg for all configurations.
a smaller Φout and less Phe loss. Still, in all configurations the Phe loss remained below 0.5%. The recovery was comparable for all configurations, with small variations attributed to different aq values and times required to reach steady state (not CPhe,b shown). The recycles resulted in accumulation of Tyr in the fermenter and of glucose and Tyr in the water removal unit aq aq and CTyr,c in Table 5). Nevertheand the crystallizer (see CGlu,c less, the mass fraction of these compounds remained well below their aqueous solubility (639 and 0.9 g kg-1, respectively21,3) and therefore, crystallization of these compounds is not expected. Economic Calculations. To assess the impact of these configurations on the economy of the process, a simplified cost model was included as described earlier. For this model the following assumptions and simplifications were made: (i) The plant operates 8000 h per year, that is, 160 batches. (ii) For cell retention, an ultrafiltration (UF) membrane is used. The required membrane area AUF is calculated from the value of Φf assuming a membrane flux of 25 kg m-2 h-1.22 For water removal, a reverse osmosis membrane as described earlier is used. (iii) Both membranes are replaced once a year. (iv) The VC concerns raw materials and membrane replacement. Other variable costs (e.g., utilities) are assumed comparable among the process configurations and hence are not included. (v) Vessels are the same for all configurations and therefore are not included in the PEC. The costs for the membrane units are calculated on the basis of the required membrane area. (vi) Revenues are calculated only on the basis of the Phe crystals recovered during fermentation. Note that, to facilitate the comparison among process configurations, several cost components are not included in this work. This means that the values obtained for VC, PEC, and NPV should not be taken as absolute values and hence, are designated as VC′, PEC′, and NPV′, respectively. Table 6 summarizes the prices used in this model and Table 7 provides the base values for the cost calculation. It can be seen that a larger Φr,2 results in a larger AUF, in agreement with the results from Table 5. On the other hand, a larger Φr,2 results in less glucose consumption. The glucose feed addition is aq controlled in order to ensure a CGlu,f of about 5 g kg-1; thus by
Figure 4. Sensitivity analysis: base case (clear bar), increase in glucose cost (striped bar), increase in glucose cost and decrease in membrane installation costs (black bar), decrease in membrane installation costs (dotted bar).
recycling Φc,out to the fermenter less glucose needs to be fed. aq (results In the same way, a larger Φr,2 results in higher CTyr,f not shown). However, Tyr is fed at a constant flow rate so this had no impact on the total Tyr consumption. Nevertheless, Tyr aq < 0.007 g kg-1). remained limiting in all cases (CTyr,f With the values from Tables 5-7 the NPV′ was calculated, from which a summary is given in Table 8. As expected, configuration B1 yields the highest costs associated to the membrane requirements (membrane replacement and PEC′), but the lowest costs in raw materials. The opposite holds for configuration B5: the membrane-associated costs are dramatically reduced when Φc,out is fully recycled to the water removal unit. Sensitivity Analysis. To explore the possibilities of these configurations, a sensitivity analysis was performed. The following cases were considered: (i) Increase in glucose cost (0.5 $/kg). The prices have continuously increased over the last 10 years up to 0.55 $/kg in January of 2008 (USDA, see reference in Table 6). (ii) Decrease in membrane installation costs (300 $/m2). With the development of better materials, the industrial use of membrane separation technologies is increasing.22 The resulting NPV’ values are shown in Figure 4. In all cases, the best result is obtained with configuration B5. These results indicate that for this system, the membrane-associated costs determine the configuration although the glucose costs are the main costs. Conclusions In this work we evaluated conceptually the recovery of Phe anhydrate crystals during a fermentation process. It was shown that crystallization is a good alternative for increasing the Phe fermentation productivity; by maintaining the Phe mass fraction in the fermenter at about 17 g kg-1 the fermentation productivity
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increased from 0.66 g kg-1 h-1, when no product removal is implemented, to 1.07 g kg-1 h-1. Product recovery by crystallization offers the additional potential advantage of obtaining the product in its commercial form, hence simplifying the downstream processing. Product recovery consisted of water removal by reverse osmosis and crystallization. Additionally, several process configurations are possible by recycling the mother liquor either to the fermenter, the water removal unit, or both. These process configurations were compared in terms of process performance and costs by means of a simplified economic model. Recycling the mother liquor to the fermenter allows a better utilization of the raw materials, while recycling it to the water removal unit drastically simplifies this process unit. It was shown that the sizing of the membranes required for cell retention and water removal has much more impact on the economic performance of the process than the consumption of raw materials, and therefore, recycling the mother liquor to the water removal unit is favored. This approach can be applied for other processes involving product recovery by crystallization during fermentation. Those processes might allow for additional configurations, like including a second crystallization operating at a lower temperature. For the Phe system considered in this study this was not an option due to the limited window of operation for Phe anhydrate crystallization and the limited temperature-dependence of Phe solubility. It also can be expected that recycling the mother liquor to the fermenter will be the preferred option for processes involving expensive raw materials and/or very low product solubility. In such cases, other effects (e.g., increased probability of contamination) should be evaluated. Furthermore, at a larger scale of operation it might be preferred to perform water removal by means of evaporation. Well-established technologies such as multieffect evaporation and vapor recompression can be applied, and the economic model should be adapted accordingly. In such a case the membrane-associated costs will be reduced, but other costs (e.g., utilities, evaporator, and compressor costs) might play a significant role. Nomenclature A ) membrane permeability (kg m-2 s-1 kPa-1) A ) membrane area (m2) B ) mass transfer coefficient (kg m-2 s-1) C ) mass fraction (g g-1, g kg-1) F ) concentration factor (unitless) g ) order of crystal growth (unitless) J ) flux (kg m-2 s-1) K ) constant (g g-1) k ) mass transfer coefficient (kg m-2 s-1) k ) constant (m s-1) L ) length (m) M ) mass (g, kg) p ) productivity (kg s-1) P ) pressure (kPa) q ) specific rate (g g-1h-1) R ) universal gas constant (J mol-1 K-1) R ) retention (unitless) T ) temperature (K) W ) molecular weight (g mol-1) Yx/y ) yield of x on y (g g-1) Greek Symbols µ ) specific cell growth rate (h-1) F ) density (g L-1) Φ ) mass flow rate (kg s-1, kg h-1) Π ) osmotic pressure (kPa)
Superscripts aq ) aqueous s ) solid, crystals Subscripts 0 ) initial b ) in the buffer vessel c ) in the crystallizer eq ) at equilibrium f ) in the fermenter g ) growth Glu ) glucose I ) inhibition m ) at the membrane m ) maintenance max ) maximal p ) in the permeate Phe ) L-phenylalanine r ) in the recycle RO ) reverse osmosis TM ) transmembrane Tyr ) L-tyrosine UF ) ultrafiltration w ) water X ) cells
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ReceiVed for reView June 22, 2009 ReVised manuscript receiVed September 10, 2009 Accepted November 10, 2009 IE901007G
